Method for producing maleic anhydride

ABSTRACT

A process for preparing maleic anhydride by heterogeneously catalyzed gas-phase oxidation of an n-butene-containing hydrocarbon stream by means of oxygen-containing gases in a shell-and-tube reactor having two successive reaction zones, where the first, feed-side reaction zone contains at least one catalyst which is suitable for the oxydehydrogenation of n-butenes to 1,3-butadiene and the second, product-side reaction zone contains at least one catalyst which is suitable for the oxidation of 1,3-butadiene to maleic anhydride, is carried out using a shell-and-tube reactor which has at least one heat transfer medium circuit in the region of the first, feed-side reaction zone and at least one further heat transfer medium circuit in the region of the second, product-side reaction zone.

Preparation of Maleic Anhydride

The present invention relates to a process for preparing maleicanhydride by heterogeneously catalyzed gas-phase oxidation of ann-butene-containing hydrocarbon stream by means of oxygen-containinggases in a shell-and-tube reactor having two successive reaction zones,where the first, feed-side reaction zone contains at least one catalystwhich is suitable for the oxydehydrogenation of n-butenes to1,3-butadiene and the second, product-side reaction zone contains atleast one catalyst which is suitable for the oxidation of 1,3-butadieneto maleic anhydride.

Maleic anhydride is an important intermediate in the synthesis ofγ-butyrolactone, tetrahydrofuran and 1,4-butanediol, which are in turnused as solvents or are processed further, for example to producepolymers such as polytetrahydrofuran or polyvinylpyrrolidone.

The preparation of maleic anhydride by single-stage heterogeneouslycatalyzed gas-phase oxidation of n-butenes or n-butane by means ofoxygen in the presence of a vanadium-, phosphorus- and oxygen-containingcatalyst is generally known and is described, for example, in Ullmann'sEncyclopedia of Industrial Chemistry, 6^(th) edition, 1999 ElectronicRelease, Chapter “MALEIC AND FUMARIC ACID—Maleic Anhydride”. In the caseof the gas-phase oxidation of n-butane using air, a maleic anhydrideselectivity of about 60% is achieved at an n-butane conversion of 85%(cf. E. Bordes, Catal. Today 16, 1993, pages 27 to 38).

A disadvantage of the use of n-butane is its often unsatisfactoryavailability at the location of the plant and the logistic problemsresulting therefrom. For this reason, n-butenes which are, for example,present in ample quantities in the C₄ fraction from a steamcracker arealso of interest as starting materials for the economical production ofmaleic anhydride. In the abovementioned gas-phase oxidation of n-butenesover the stated vanadium-, phosphorus- and oxygen-containing catalysts,the yield of maleic anhydride is restricted to values of about 50 mol %.E. Bordes reports an achievable maleic anhydride selectivity of 50 mol %at a 1-butene conversion of 95% (cf. E. Bordes, Catal. Today 16, 1993,pages 27 to 38). In addition, the selective oxidation of n-butenes tomaleic anhydride forms a series of undesirable by-products compared tothe selective oxidation of n-butane.

A further starting material known for the preparation of maleicanhydride is 1,3-butadiene. Thus, DE-A 2 241 918 describes the gas-phaseoxidation of 1,3-butadiene to maleic anhydride in the presence ofmolybdenum-, antimony- and oxygen-containing catalysts. Disadvantages ofthe direct use of pure 1,3-butadiene are its unsatisfactory availabilityand its high price.

An inexpensive possibility for preparing a 1,3-butadiene-containingstream is the oxydehydrogenation of n-butenes or n-butene-containingstreams. Thus, A. Yoshioka et al. in Hydrocarbon Processing, 63,November 1984, pages 97 to 100, describe the oxydehydrogenation of araffinate II stream to give a 1,3-butadiene yield of 78%.

DE-A 26 03 770 describes the preparation of maleic anhydride by reactionof n-butenes with air in a fluidized-bed reactor in the presence of anoxidation catalyst which is effective in the oxydehydrogenation ofn-butenes to 1,3-butadiene and an oxidation catalyst which is effectivein the oxidation of 1,3-butadiene to maleic anhydride. A disadvantage ofthis process is a low maleic anhydride yield of 30%.

DE-A 25 39 106 teaches the preparation of maleic anhydride by reactionof n-butane and n-butene-containing streams with air in a shell-and-tubereactor having an undivided, isothermal reaction zone containing twodifferent catalysts. The catalyst present on the feed sideoxydehydrogenates n-butane and n-butene to 1,3-butadiene. The catalystpresent on the product side oxidizes the 1,3-butadiene formed to maleicanhydride. A maleic anhydride yield of 60% was achieved using a cis- andtrans-2-butene mixture and an oxydehydrogenation catalyst based onbismuth molybdate and an oxidation catalyst based on antimony molybdate.

According to the present invention, it has been recognized that theabove-described direct connection of oxydehydrogenation and oxidation inseries and, in particular, isothermal operation of both processesresults in many disadvantages. Furthermore, it has been recognized,according to the present invention, that very different reactiontemperatures are required for optimal operation of theoxydehydrogenation of n-butenes to 1,3-butadiene and of the oxidation of1,3-butadiene to maleic anhydride. In addition, the amounts of heatevolved in the two reaction steps are very different. While only about130 kJ/mol are liberated in the oxydehydrogenation of n-butenes to1,3-butadiene, the oxidation of 1,3-butadiene to maleic anhydridereleases about 990 kJ/mol. Heat removal matched to the evolution of heatis thus not possible in reality. Furthermore, matching to the optimumoperating conditions of the two catalysts is likewise not possible orpossible only to a very limited extent. Thus, effects due to differentcatalyst activities, to deactivation processes proceeding at differentrates, to fluctuations in the purity and quality of the feed stream orto load changes (changes in the flow velocity and/or the feed rate) canbe corrected only to a very restrictive extent, if at all.

It is an object of the present invention to develop a process forpreparing maleic anhydride by heterogeneously catalyzed gas-phaseoxidation of an inexpensive and readily available hydrocarbon stream bymeans of oxygen, which process does not have the abovementioneddisadvantages and makes possible a high conversion, a high selectivityand a high yield of desired product at a high hydrocarbon throughputover the catalyst and thus gives a high space-time yield. A furtherobject of the present invention is to allow flexible operation whichmakes it possible to achieve a high space-time yield over a long periodof time even in the case of fluctuations in the amount, quality orpurity of the starting materials or in the event of progressive catalystdeactivation.

We have found that this object is achieved by a process for preparingmaleic anhydride by heterogeneously catalyzed gas-phase oxidation of ann-butene-containing hydrocarbon stream by means of oxygen-containinggases in a shell-and-tube reactor having two successive reaction zones,where the first, feed-side reaction zone contains at least one catalystwhich is suitable for the oxydehydrogenation of n-butenes to1,3-butadiene and the second, product-side reaction zone contains atleast one catalyst which is suitable for the oxidation of 1,3-butadieneto maleic anhydride, wherein the shell-and-tube reactor has at least oneheat transfer medium circuit in the region of the first, feed-sidereaction zone and at least one further heat transfer medium circuit inthe region of the second, product-side reaction zone.

For the purposes of the present invention, the term “shell-and-tubereactor” refers to a reactor which contains at least one reaction tubewhich is surrounded by a heat transfer medium for the purpose of heatingand/or cooling. In general, shell-and-tube reactors used industriallycontain from a few hundred to tens of thousands of reactor tubesconnected in parallel. If a number of individual shell-and-tube reactors(in the sense of shell-and-tube reaction apparatuses) are connected inparallel, these should be regarded as the equivalent of oneshell-and-tube reactor and are hereinafter encompassed by the termshell-and-tube reactor.

For the purposes of the present invention, the term first, feed-sidereaction zone is the region within the shell-and-tube reactor whichcontains at least one catalyst suitable for the oxydehydrogenation ofn-butenes to 1,3-butadiene. The term second, product-side reaction zoneis the region within the shell-and-tube reactor which contains at leastone catalyst suitable for the oxidation of 1,3-butadiene to maleicanhydride.

In the process of the present invention, it is essential to use ashell-and-tube reactor which has one heat transfer medium circuit in theregion of the first, feed-side reaction zone and a further heat transfermedium circuit in the region of the second, product-side reaction zone.Thus, for example, each of the two reaction zones can be heated/cooledby means of one, two or three or more heat transfer medium circuits.Preference is given to using a shell-and-tube reactor which hasprecisely one heat transfer medium circuit in the region of the firstreaction zone and precisely one heat transfer medium circuit in theregion of the second reaction zone.

Shell-and-tube reactors used in the process of the present invention canin principle be any known shell-and-tube reactors having two or moreheat transfer medium circuits. Preference is given to usingshell-and-tube reactors having two heat transfer medium circuits.Suitable shell-and-tube reactors are described, for example, in U.S.Pat. No. 3,147,084, DE-C 28 30 765, EP-A 0 911 313 and EP-A 1 097 745.

In general, the shell-and-tube reactors which can be used in the processof the present invention comprise an outer reactor body havingconnections for the introduction and discharge of the reaction gases atthe top and bottom. The upper and lower ends of the reactor tubes arewelded into upper and lower tube plates. The space surrounding thereactor tubes between the upper and lower tube plates is generallydivided into two zones by means of a dividing plate. Each of the twozones usually has two connections on the reactor body for theintroduction and discharge of the heat transfer medium. In general, thedividing plate is located at the height of the transition region betweenthe two successive reaction zones.

In one embodiment, the dividing plate can, as described in U.S. Pat. No.3,147,084, be fixed to the reactor tubes and seal the two heat transfermedium zones from one another.

In another embodiment, the dividing plate can have a very small gapbetween it and the individual reactor tubes so as to allow unhinderedthermal expansion of the reactor tubes and thus prevent mechanicalstresses and consequent corrosion and mechanical damage. Such aconstruction is described, for example, in DE-C 28 30 765. In a furtherembodiment, the reactor tubes are widened in the direction of thedividing plate in the region of the dividing plate in order to achieve afurther reduction in the gap between the tube and the reaction plate.The hole through the dividing plate can be provided on its inside withadditional corrugations to improve the seal between the two heatexchange media, as described in EP-A 1 097745. It is also possible forthe dividing plate to be fixed to the reactor tubes but to have anannular gap between it and the reactor body. To avoid mechanicalstresses due to nonuniform thermal expansion, the reactor body can alsobe equipped with expansion zones, for example in the form ofsemicircular recesses. A reactor constructed in such a way is described,for example, in EP-A 1 097 745.

The reactor tubes in the abovementioned shell-and-tube reactors areusually made of ferritic steel and typically have a wall thickness offrom 1 to 3 mm. Their internal diameter is generally from 15 to 30 mm.The number of reactor tubes per shell-and-tube reactor is usually in therange from 5000 to 35 000, although a number above 35 000 can also beemployed in particularly large plants. The reactor tubes are normallydistributed uniformly within the reactor body.

The heat transfer medium zones may also contain various internals fordirecting the flow of the heat transfer medium. Preference is given tointernals which deflect at least part of the circulating heat transfermedium in a radial direction, i.e. within the cross section, and thusallow largely perpendicular, i.e. radial, flow onto the reactor tubesrunning in the axial direction. Preference is given to internals whichpromote essentially meandering flow of at least part of the heattransfer medium. Explicit mention may be made of an alternativearrangement of “donut-shaped” deflecting plates having a central holethrough which fluid can flow and circular deflection plates having a gapfor the flow of fluid between the reactor body and the deflection plate,as described, for example, in EP-A 1 097 745.

Suitable heat exchange media are, in particular, fluid cooling/heatingmedia. The use of salt melts, e.g. potassium nitrate, potassium nitrite,sodium nitrate and/or sodium nitrite, or of low-melting metals such assodium or alloys of various metals is particularly useful. Preference isgiven to using salt melts.

The direction of flow of the heat transfer medium in the reactor can inprinciple be either from the top downward or from the bottom upward andcan be chosen independently for the two reaction zones. For hydrodynamicreasons, especially because of the induced convection, flow of the heattransfer medium from the bottom upward in the reactor is preferred inthe process of the present invention.

The reaction gas can flow through the shell-and-tube reactors which canbe used in the process of the present invention either from the topdownward or from the bottom upward. In the first-named case, thecatalyst which is suitable for the oxydehydrogenation of n-butenes to1,3-butadiene is present in the upper reaction zone and the catalystwhich is suitable for the oxidation of 1,3-butadiene to maleic anhydrideis located in the lower reaction zone. In the second case, the catalystis arranged in the converse manner.

If the heat transfer medium flows in the same direction as the reactiongas, this is referred to as cocurrent operation; if the heat transfermedium flows in the opposite direction, this is referred to ascountercurrent operation. In the process of the present invention,preference is given to cocurrent operation, since this generally allowsbetter distribution of the hot spot.

On the basis of the preferred flow of heat transfer medium from thebottom upward and a countercurrent mode of operation, upward flow of thereaction gases and the two heat exchange media is particularly preferredin the process of the present invention.

Furthermore, the shell-and-tube reactor to be used can also contain oneor more integrated or upstream preheating zones which heat the gasmixture entering the reactor. A preheating zone integrated into ashell-and-tube reactor can be achieved, for example, by means of reactortubes which are filled with inert material and may be surrounded by aheat transfer medium. As inert material, it is in principle possible touse all materials which do not contribute to the chemical reaction ofthe reaction gas flowing through them, i.e. do not induce or catalyze aheterogeneously catalyzed reaction, and which result in a maximumpressure drop below the maximum value tolerable in the case of thespecific plant. Suitable materials are, for example, oxidic materials-,e.g. Al₂O₃, SiC, or metallic materials, e.g. stainless steel. The inertmaterial can be present, for example, in the form of shaped bodies,meshes, open sponges and knitteds or internals as are customarily alsoused in static mixers. Preference is given to shaped bodies such asspheres, pellets, hollow cylinders, rings, trilobes, tristars, wagonwheels, extrudates or irregular, crushed bodies.

In the process of the present invention, the catalysts of the twosuccessive reaction zones can, for example, follow one another directlyor be separated from one another by an empty space or an intermediatebed. The transition between the two catalysts, the empty space or theintermediate bed is in each case preferably at the approximate height ofor in the vicinity of the dividing plate.

In the process of the present invention, the reaction gas from thefirst, feed-side reaction zone is preferably passed through an inertintermediate bed before it enters the second, product-side reactionzone. Inert materials suitable for this intermediate bed are generallymaterials which do not, under the prevailing reaction conditions,contribute significantly to chemical reaction of the reaction gasflowing through the bed, i.e. do not induce or catalyze aheterogeneously catalyzed reaction, and which result in a maximumpressure drop below the maximum which can be tolerated in the specificplant. Examples of suitable materials are oxidic materials, e.g. Al₂O₃,SiC, or metallic materials, e.g. stainless steel. The inert material canbe present, for example, in the form of shaped bodies, meshes, opensponges and knitteds or internals as are customarily also used in staticmixers. Preference is given to shaped bodies such as spheres, pellets,hollow cylinders, rings, trilobes, tristars, wagon wheels, extrudates orirregular, crushed bodies. The maximum diameter of the preferred shapedbodies is preferably from 1/10 to at most 1/1 of the internal diameterof the reactor tubes.

Depending on whether the second reaction zone is operated at a higher orlower temperature, the inert intermediate layer makes heating or coolingof the reaction gas possible. To enable this to be achieved, the majorpart of the inert intermediate bed is generally in the region of thesecond, hotter or colder zone. Furthermore, the inert intermediate bedallows precipitation of relatively high molecular weight by-productsfrom the first reaction zone and counters deposition on the catalyst ofthe second reaction zone and thus a gradual increase in the pressuredrop due to such deposits on the catalyst.

The length of the intermediate bed is preferably such that the majorpart of the temperature difference between the two reaction zones occurswithin the intermediate bed. The length of the intermediate bed may alsoneed to be chosen according to the amount of relatively high molecularweight by-products formed in the first reaction zone and the desireddegree of removal of these by-products. These parameters can bedetermined, for example, by means of simple laboratory or pilot planttests.

The inert material of the inert intermediate bed preferably has an emptyvolume fraction of from 40 to 99.5%, with the empty volume fractionbeing defined as: empty volume fraction=[(total volume)−(geometricvolume)]/(total volume). Here, the “total volume” is the total volume ofthe inert intermediate bed in the reaction tube. The “geometric volume”is the macroscopic volume of the inert solid material including anypores present within the inert material. At an empty volume fractionbelow 40%, the pressure drop generally increases significantly. At anempty volume fraction greater than 99.5%, the desired effect of heatingor cooling and the precipitation of relatively high molecular weightby-products is generally unsatisfactory. The empty volume fraction ofthe inert material is preferably from 45 to 99%.

Catalysts which can be used in the process of the present invention arein principle all catalysts suitable for the oxydehydrogenation ofn-butenes to 1,3-butadiene and all catalysts suitable for the oxidationof 1,3-butadiene to maleic anhydride.

Particularly useful catalysts for the oxydehydrogenation of n-butenes to1,3-butadiene are generally based on an Mo—Bi—O-containing multimetaloxide system which generally further comprises iron. In general, thecatalyst system further comprises additional components from groups 1 to15 of the Periodic Table, for example potassium, magnesium, zirconium,chromium, nickel, cobalt, cadmium, tin, lead, germanium, lanthanum,manganese, tungsten, phosphorus, aluminum or silicon.

Suitable catalysts and their production are described, for example, inU.S. Pat. No. 4,423,281 (Mo₁₂BiNi₈Pb_(0.5)Cr₃K_(0.2)O_(x) andMo₁₂Bi_(b)Ni₇Al₃Cr_(0.5)K_(0.5)O_(x)), U.S. Pat. No. 4,336,409(Mo₁₂BiNi₆Cd₂Cr₃P_(0.5)O_(x)), DE-A 26 00 128(Mo₁₂BiNi_(0.5)Cr₃P_(0.5)Mg_(7.5)K_(0.1)O_(x)+SiO₂) and DE-A 24 40 329(Mo₁₂BiCo_(4.5)Ni_(2.5)Cr₃P_(0.5)K_(0.1)O_(x)), which are explicitlyincorporated by reference.

The stoichiometry of the active composition of many of the multimetaloxide catalysts suitable for the oxydehydrogenation of n-butenes to1,3-butadiene can be described by the formula (I)Mo₁₂Bi_(a)Fe_(b)Co_(c)Ni_(d)Cr_(e)X¹ _(f)K_(g)O_(x)  (I),where the variables have the following meanings:

-   X¹=W, Sn, Mn, La, Ce, Ge, Ti, Zr, Hf, Nb, P, Si, Sb, Al, Cd and/or    Mg;-   a=0.5 to 5, preferably from 0.5 to 2;-   b=0 to 5, preferably from 2 to 4;-   c=0 to 10, preferably from 3 to 10;-   d=0 to 10;-   e=0 to 10, preferably from 0.1 to 4;-   f=0 to 5, preferably from 0.1 to 2;-   g=0 to 2, preferably from 0.01 to 1; and-   x=a number determined by the valence and abundance of the elements    other than oxygen in (I).

In the process of the present invention, the oxydehydrogenation ispreferably carried out using an Mo—Bi—Fe—O-containing multimetal oxidesystem, with an Mo—Bi—Fe—Cr—O- or Mo—Bi—Fe—Zr—O-containing multimetaloxide system being particularly preferred. Preferred systems aredescribed, for example, in U.S. Pat. No. 4,547,615(Mo₁₂BiFe_(0.1)Ni₈ZrCr₃K_(0.2)O_(x) andMo₁₂BiFe_(0.1)Ni₈AlCr₃K_(0.2)O_(x)), U.S. Pat. No. 4,424,141(Mo₁₂BiFe₃Co_(4.5)Ni_(2.5)P_(0.5)K_(0.1)O_(x)+SiO₂), DE-A 25 30 959(Mo₁₂BiFe₃Co_(4.5)Ni_(2.5)Cr_(0.5)K_(0.1)O_(x),Mo_(13.75)BiFe₃Co_(4.5)Ni_(2.5)Ge_(0.5)K_(0.8)O_(x),Mo₁₂BiFe₃Co_(4.5)Ni_(2.5)Mn_(0.5)K_(0.1)O_(x) andMo₁₂BiFe₃Co_(4.5)Ni_(2.5)La_(0.5)K_(0.1)O_(x)), U.S. Pat. No. 3,911,039(Mo₁₂BiFe₃Co_(4.5)Ni_(2.5)Sn_(0.5)K_(0.1)O_(x)), DE-A 25 30 959 and DE-A24 47 825 (Mo₁₂BiFe₃Co_(4.5)Ni_(2.5)W_(0.5)K_(0.1)O_(x)). Production andcharacterization of the catalysts mentioned are comprehensivelydescribed in the cited documents, which are hereby explicitlyincorporated by reference.

The catalysts used for the oxidation of 1,3-butadiene to maleicanhydride are generally based on a multimetal oxide system comprisingmolybdenum oxide and preferably on an essentially Mo—Sb—O-containing orMo—V—O-containing multimetal oxide system.

In the case of the Mo—Sb—O-containing multimetal oxide system, thecatalyst system generally further comprises additional components fromgroups 1 to 8 and 14 of the Periodic Table and from the lanthanideseries, for example lithium, titanium, vanadium, niobium, tungsten,iron, tin and cerium.

In the case of the Mo—Sb—O-containing multimetal oxide system, theactive composition of many of the multimetal oxide catalysts suitablefor the oxidation of 1,3-butadiene to maleic anhydride has the formula(II)SbMo_(a)Sn_(b)X² _(c)O_(x)  (II),where the variables have the following meanings:

-   X²=V, Fe, Ni, Li, Ce, Nb, Ta, W, Ti, Zr, B, P, Al and/or Mg;-   a=1 to 10, preferably from 2 to 4;-   b=0 to 5, preferably from 0.01 to 2;-   c=0 to 3, preferably from 0.01 to 1.5; and-   x=a number determined by the valence and abundance of the elements    other than oxygen in (II).

Suitable catalysts and their production are described, for example, inJP-A 05 057 188 (Sb₂Mo₁₀O_(x) corresponding to SbMo₅O_(x)), DE-A 22 41918 (Sb₄Mo_(6.12)O_(x) corresponding to SbMo_(1.53)O_(x)), DE-A 23 22186 (SbMo₃V_(0.1)Fe_(0.2)W_(0.06)O_(x)), DE 27 02 606(SbMo₃V_(0.1)Li_(0.1)W_(0.06)O_(x) andSbMo₃V_(0.1)Ce_(0.1)W_(0.06)O_(x)) and DE-A 28 13 424(SbMo_(3.06)T_(0.6)Nb_(0.1)O_(x)). The production and characterizationof the catalysts mentioned are comprehensively described in the citeddocuments, which are hereby explicitly incorporated by reference.

In the case of the Mo—V—O-containing multimetal oxide, the activecomposition of many of the multimetal oxide catalysts suitable for theoxidation of 1,3-butadiene to maleic anhydride has the formula (III)Mo₁₂V_(a)W_(b)X³ _(c)O_(x)  (III),where the variables have the following meanings:

-   X³=La, Mn, Fe, Cu, Al, Co, Ni, Bi, Ag, P, Zn, Cd, As, Cr, Sn, U, Ti,    Nb, Ge, alkali metal and/or alkaline earth metal;-   a=0.1 to 12, preferably from 1.5 to 10;-   b=0 to 5, preferably from 0.1 to 4;-   c=0 to 12, preferably from 0.1 to 10; and-   x=a number determined by the valence and abundance of the elements    other than oxygen in (III).

Suitable catalysts and their production are described, for example, inU.S. Pat. No. 3,893,951 (Mo₁₂V₃W_(1.2)O_(x), Mo₁₂V₃W_(1.2)Sn₆O_(x)),U.S. Pat. No. 4,157,987 (Mo₁₂V₃W_(1.2)Ce₃O_(x), Mo₁₂V₃W_(1.2)Ce₂CoO_(x),Mo₁₂V₃W_(1.2)Ce₂Cu₂O_(x)), DE-A 24 59 092 (Mo₁₂V₃W_(1.2)U₂O_(x)), U.S.Pat. No. 4,170,570 (Mo₁₂V₃W_(1.2)Cu₂Sn_(0.5)O_(x)), U.S. Pat. No.4,378,309 (Mo₁₂V₃Cu_(0.5)GeO_(x)), U.S. Pat. No. 4,042,533(Mo₁₂V₃W_(1.2)Ti_(0.5)O_(x)), U.S. Pat. No. 4,115,441(Mo₁₂V₃GeFe_(0.1)O_(x)), U.S. Pat. No. 4,138,366(Mo₁₂V₃SbCd_(0.2)P_(0.1)O_(x)), and U.S. Pat. No. 4,250,054(Mo₁₂V₃W_(1.2)La_(0.5)Co_(0.1)O_(x)). Production and characterization ofthe catalysts mentioned are comprehensively described in the citeddocuments, which are hereby explicitly incorporated by reference.

The two catalysts for oxydehydrogenation and for oxidation are generallyused as shaped bodies having a mean size of above 2 mm. Owing to thepressure drop, of which note has to be taken, when carrying out theprocess, relatively small shaped bodies are generally unsuitable.Suitable shaped bodies which may be mentioned are, for example, pellets,cylinders, hollow cylinders, rings, spheres, rods, wagon wheels orextrudates. Particular shapes such as “trilobes” and “tristars” (cf.EP-A-0 593 646) or shaped bodies having at least one constriction on theoutside (cf. U.S. Pat. No. 5,168,090) are likewise possible.

In general, the catalysts to be used can be employed as all-activecatalysts. In this case, the entire catalyst body consists of the activecomposition, including any auxiliaries such as graphite or pore formers,together with further components. In particular, theMo—Bi—Fe—O-containing catalyst which is preferably used for theoxydehydrogenation of n-butenes to 1,3-butadiene is advantageously usedas an all-active catalyst. It is also possible to apply the activecompositions of the catalysts to a support, for example an inorganic,oxidic shaped body. Such catalysts are generally referred to as coatedcatalysts. In particular, the Mo—Sb—O- or Mo—V—O-containing catalystwhich is preferably used for the oxidation of the 1,3-butadienegenerated in the first step to maleic anhydride is advantageously usedas a coated catalyst.

In the process of the present invention, the catalysts used in the tworeaction zones are different catalysts, one of which is optimized forthe oxydehydrogenation of n-butenes to 1,3-butadiene and the other isoptimized for the oxidation of 1,3-butadiene to maleic anhydride.Likewise, the optimized reaction temperatures are set in each of the tworeaction zones.

As regards the way the catalysts are installed in the first, feed-sidereaction zone and in the second, product-side reaction zone in theprocess of the present invention, a number of variants are possible. Inone variant, each reaction zone contains a homogeneous catalyst bed,i.e. a catalyst bed which has the same average composition and the sameaverage activity per unit volume throughout the respective reactionzone. A catalyst bed can be made up of shaped bodies of the samecatalyst, of shaped bodies of a mixture of various catalysts or ofshaped bodies (same catalyst or mixture of various catalysts) which aremixed, i.e. “diluted”, with an inert material. Suitable inert materialsare in principle all shaped bodies which are also suitable for use inthe preheating zones or as inert intermediate bed. Reference may be madeto the information provided above.

In another variant, the first, feed-side reaction zone and/or thesecond, product-side reaction zone contain/contains a heterogeneouscatalyst bed, i.e. a catalyst bed which has a composition and anactivity per unit volume which vary over the length of the respectivereaction zone. A heterogeneous catalyst bed is generally achieved bymeans of a heterogeneous mixture of various catalysts or shaped bodies(same catalyst or mixture of various catalysts) which areheterogeneously mixed, i.e. “diluted”, with an inert material. The localactivity within the reaction zone is thus set via the composition of themixture. When using a heterogeneous catalyst bed, an increase in theactivity in the flow direction of the reaction gas, i.e. from the feedside to the product side, is generally advantageous. The activity canincrease uniformly in the flow direction of the reaction gas or canincrease stepwise at one or more points.

In general, the optimized reaction temperature for theoxydehydrogenation of n-butenes to 1,3-butadiene is below that for theoxidation of 1,3-butadiene to maleic anhydride. This appliesparticularly when the catalyst used for the oxidation of 1,3-butadieneto maleic anhydride has the basic composition Mo—Sb—O. In the case of acorresponding Mo—V—O-based catalyst, specific compositions of thiscatalyst can result in the optimized reaction temperature for theoxydehydrogenation of n-butenes to 1,3-butadiene being above that forthe oxidation of 1,3-butadiene to maleic anhydride.

In the present context, the reaction temperature is the temperature ofthe catalyst bed present in this reaction zone which would prevail ifthe process were carried out in the absence of a chemical reaction. Ifthis temperature is not exactly the same at all points, the term refersto the arithmetic mean of the temperatures along the reaction zone. Itthus corresponds approximately to the temperature of the surroundingheat transfer medium.

In the case of a catalyst having the basic composition Mo—Sb—O which issuitable for the oxidation of 1,3-butadiene to maleic anhydride in thesecond, product-side reaction zone, this is preferably operated at areaction temperature which is at least 10° C. higher, particularlypreferably at least 30° C. higher, than that of the catalyst which issuitable for the oxydehydrogenation of n-butenes to 1,3-butadiene in thefirst, feed-side reaction zone.

It is a specific feature of the process of the present invention that,depending on requirements, for example due to changes in the throughputover the catalyst or due to ageing of the catalysts, the reactiontemperatures of the two reaction zones can be regulated independently.

The oxydehydrogenation of n-butenes to 1,3-butadiene in the first,feed-side reaction zone in the process of the present invention isgenerally carried out at from 220 to 490° C., preferably from 250 to450° C.

The oxidation of 1,3-butadiene to maleic anhydride in the second,product-side reaction zone in the process of the present invention isgenerally carried out at from 190 to 500° C., preferably from 230 to460° C.

As regards the reaction pressure, there are generally no particularrequirements in the process of the present invention. For practicalreasons, it is normal to select a pressure at the inlet of the reactorwhich is sufficient to overcome the flow resistances present in theplant and in the subsequent work-up. This reactor inlet pressure isgenerally from 0.005 to 1 MPa gauge pressure, preferably from 0.01 to0.5 MPa gauge pressure. The gas pressure employed at the inlet region ofthe reactor naturally drops significantly over the bed of catalysts andinert sections.

n-Butene-containing hydrocarbon streams used in the process of thepresent invention are generally hydrocarbon streams which have a totaln-butene content (1-butene, 2-trans-butene and 2-cis-butene) of ≧10% byweight, preferably ≧30% by weight. The n-butene-containing hydrocarbonstream may further comprise aliphatic and aromatic, saturated andunsaturated hydrocarbons. Examples which may be mentioned are methane,ethane, ethene, propane, propene, 1,3-butadiene, n-butane, isobutane,isobutene, 1,3-pentadiene, 1,4-pentadiene, 1-pentene, 2-cis-pentene,2-trans-pentene, n-pentane, cyclopentadiene, dicyclopentadiene,cyclopentene, cyclopentane, hexenes, hexanes, cyclohexane and benzene.

The n-butene-containing hydrocarbon streams used in the process of thepresent invention are preferably streams from natural gas, steamcrackers or FCC plants. Particular preference is given to raffinate IIwhich is obtained from the crude C₄ fraction from steam crackers.Raffinate II is the C₄ stream from which 1,3-butadiene has been largelyremoved or converted into n-butenes by selective hydrogenation andisobutene has been removed. Its n-butene content is generally from 50 to95% by weight. A typical but nonlimiting composition of a raffinate IIstream is given in Table 1:

TABLE 1 Typical composition of a raffinate II stream. Raffinate II1-butene from 20 to 60% by weight 2-trans-butene from 10 to 30% byweight 2-cis-butene from 5 to 20% by weight n-butane from 5 to 35% byweight isobutane from 1 to 10% by weight isobutene from 1 to 2% byweight

The addition of the hydrocarbons is generally quantity-regulated, i.e.regulated so that a defined amount is added per unit time. Thehydrocarbon can be metered in in liquid or gaseous form. Preference isgiven to metering in the hydrocarbon in liquid form and subsequentlyvaporizing it before entry into the shell-and-tube reactor.

The oxygen-containing gas used in the process of the present inventionis generally air, synthetic air, an oxygen-enriched gas or “pure”oxygen, for example oxygen from the fractionation of air. Theoxygen-containing gas, too, is added in a quantity-regulated manner.

The gas to be passed through the two reaction zones generally comprisesinert gas. The proportion of inert gas at the beginning is usually from30 to 95% by volume. Inert gases are all gases which do not contributedirectly to the formation of maleic anhydride, for example nitrogen,steam, noble gases, carbon monoxide, carbon dioxide, oxygenated andnonoxygenated hydrocarbons containing less than four carbon atoms, e.g.methane, ethane, propane, methanol, formaldehyde, formic acid, ethanol,acetaldehyde, acetic acid, propanol, propionaldehyde, propionic acid,acrolein, crotonaldehyde, acrylic acid) and mixtures thereof. Ingeneral, the inert gas is introduced into the system via theoxygen-containing gas. However, it is also possible for further inertgases to be fed in separately. Thus, for example, the process of thepresent invention can be carried out with introduction of generally upto 50% by weight of steam. The process of the present invention ispreferably carried out without introduction of additional inert gases.

The process of the present invention is generally carried out in a“single pass”. The reaction gas used generally comprises from 0.1 to 30%by volume, preferably from 0.5 to 7% by volume, of n-butenes and from 5to 50% by volume, preferably from 10 to 40% by volume, of oxygen. Theconversion of n-butenes is generally from 50 to 100%, preferably from 75to 100% and particularly preferably from 85 to 100%. The unreactedn-butenes can be separated from the maleic anhydride formed and furtherproducts and then recirculated to the initial feed. Preference is givento operating the process in a single pass with a very high n-buteneconversion and no recirculation of unreacted n-butenes.

The maleic anhydride can be separated off by, for example, absorption ina suitable absorption medium. Suitable absorption media are, forexample, water or organic solvents. In the case of absorption in water,maleic anhydride is hydrated to form maleic acid. Preference is given toabsorption in an organic solvent. Suitable organic solvents are, forexample, the high-boiling solvents mentioned in WO 97/43242, for exampletricresyl phosphate, dibutyl maleate, high molecular weight wax,aromatic hydrocarbons having a boiling point above 140° C. ordi-C₄–C₈-alkyl phthalates such as dibutyl phthalate. Oxygenatedhydrocarbon by-products are generally also absorbed in the solventsmentioned. The absorption can be carried out, for example, at from 60 to160° C. and a pressure of from 0.1 to 0.5 MPa abs or above. Suitablemethods are, for instance, passing the gaseous, cooled or uncooledreactor output through a vessel filled with absorption liquid orspraying the absorption liquid into the gas stream. Appropriate methodsof scrubbing gas streams are known to those skilled in the art.

Furthermore, we have found the use of a shell-and-tube reactor havingtwo successive reaction zones and at least one heat transfer mediumcircuit in the region of the first, feed-side reaction zone and at leastone further heat transfer medium circuit in the region of the second,product-side reaction zone in the heterogeneously catalyzed gas-phaseoxidation of an n-butene-containing hydrocarbon stream by means ofoxygen to form maleic anhydride.

Preference is given to the use of a shell-and-tube reactor having twosuccessive reaction zones and one heat transfer medium circuit in theregion of the first, feed-side reaction zone and a further heat transfermedium circuit in the region of the second, product-side reaction zonein the heterogeneously catalyzed gas-phase oxidation of ann-butene-containing hydrocarbon stream by means of oxygen to form maleicanhydride.

A description is given below of a preferred embodiment of thepreparation of maleic anhydride by heterogeneously catalyzed gas-phaseoxidation of an n-butene-containing hydrocarbon stream by means of anoxygen-containing gas in a shell-and-tube reactor having two successivereaction zones and having a heat transfer medium circuit in the regionof the first feed-side-reaction zone and a further heat transfer mediumcircuit in the region of the second, product-side reaction zone.

The heat transfer medium flows from the bottom upward through each ofthe two heat transfer medium circuits of the shell-and-tube reactor. Theshell-and-tube reactor is operated in cocurrent, i.e. the reaction gas(feed gas) is fed in from the bottom. The reaction gas used is a gasmixture of raffinate II in air having an n-butene content of from 0.5 to7% by volume. In the lowermost zone of the bed, which functions aspreheating zone, the shell-and-tube reactor contains inert material. Thenext zone of the bed in an upward direction, viz. the first reactionzone, contains the catalyst for the oxydehydrogenation of the n-butenesto 1,3-butadiene, which comprises an Mo—Bi—Fe—O-containing multimetaloxide system as active composition. The first reaction zone is operatedin a temperature range from 250 to 450° C. In the region of or in thevicinity of the dividing plate located in the shell-and-tube reactor toseparate the two heat transfer medium circuits, there is an inertintermediate bed having an empty volume fraction in the range from 40 to99.5%. The following, second reaction zone contains the catalyst for theoxidation of 1,3-butadiene to maleic anhydride, which comprises anMo—Sb—O-containing or Mo—V—O-containing multimetal oxide system asactive composition. The second reaction zone is operated in atemperature range from 230 to 460° C. The pressure of the reaction gasat the reactor inlet is in the range from 0.01 to 0.5 MPa gaugepressure. Temperature and space velocity over the catalyst are generallyselected so that an overall conversion of n-butenes of ≧50%, preferably≧75% and particularly preferably ≧85%, results. The reaction gas takenoff at the top of the reactor is passed to an absorption stage toseparate off the maleic anhydride formed.

The process of the present invention makes it possible to achieve a highconversion, a high selectivity, a high yield of desired product and thusa high space-time yield in the preparation of maleic anhydride byheterogeneously catalyzed gas-phase oxidation using an inexpensive andreadily available hydrocarbon stream at a high hydrocarbon throughputover the catalyst. Due to the two heat transfer medium circuits whichallow the temperature in the two reaction zones to be regulatedseparately, the process of the present invention can be carried outflexibly so as to make it possible to achieve a high space-time yieldover a long period of time even in the case of fluctuations in theamount, quality or purity of starting materials or in the event ofprogressive catalyst deactivation. Compared to a process carried out intwo shell-and-tube reactors connected in series, carrying out theprocess in a single shell-and-tube reactor having two successivereaction zones leads to considerable capital cost savings and to asignificant simplication of the overall plant.

In addition, the inert intermediate bed located between the first,feed-side reaction zone and the second, product-side reaction zone in apreferred variant of the process of the present invention makes itpossible for relatively high molecular weight by-products from the firstreaction zone, for example oligomers, polymers or carbon, to beprecipitated and counters deposition on the catalyst of the secondreaction zone and thus a gradual increase in the pressure drop caused bysuch deposits on the catalyst.

The maleic anhydride obtained can, for example, be processed further toproduce γ-butyrolactone, tetrahydrofuran, 1,4-butanediol or mixturesthereof. Suitable processes are known to those skilled in the art. Forthe sake of completeness, reference is made to the two documents WO97/43234 (direct hydrogenation of maleic anhydride in the gas phase) andWO 97/43242 (hydrogenation of a maleic diester in the gas phase).

EXAMPLES Definitions

The parameters referred to in this text are, unless indicated otherwise,defined as follows: $\begin{matrix}{{{{Conversion}\mspace{14mu} C} = \frac{n_{{HC},{reactor},{i\; n}} - n_{{HC},{reactor},{out}}}{n_{{HC},{reactor},{i\; n}}}}\mspace{11mu}} \\{\;{{{Selectivity}\mspace{14mu} S_{MA}} = \frac{n_{{MA},{reactor},{out}}}{n_{{HC},{reactor},{i\; n}} - n_{{HC},{reactor},{out}}}}}\end{matrix}$ Yield Y _(MA) =U·S _(MA)

-   C Conversion of hydrocarbons per pass through the reactor-   S_(MA) Selectivity to maleic anhydride per pass through the reactor-   Y_(MA) Yield of maleic anhydride per pass through the reactor-   n_(HC, reactor, in) Molar flow of hydrocarbons at the reactor inlet    [mol/h]-   n_(HC, reactor, out) Molar flow of hydrocarbons at the reactor    outlet [mol/h]-   n_(HC, plant, in) Molar flow of hydrocarbons at the inlet to the    plant [mol/h]-   n_(HC, plant, out) Molar flow of hydrocarbons at the outlet of the    plant [mol/h]-   n_(HC, reactor, out) Molar flow of maleic anhydride at the reactor    outlet [mol/h]-   n_(MA, plant, out) Molar flow of maleic anhydride at the outlet of    the plant [mol/h]    Production of the Oxydehydrogenation Catalyst K1

1750.9 g of aqueous cobalt nitrate solution having a free HNO₃ contentof 0.2% by weight and a Co content of 12.5% by weight (=3.71 mol of Co)were placed in a heatable 10 l stirred vessel made of glass. Whilestirring, 626.25 g of solid Fe(NO₃)₃.9H₂O having an Fe content of 14.2%by weight (=1.59 mol of Fe) were dissolved at room temperature in theinitial charge of cobalt nitrate solution. 599.5 g of bismuth nitratesolution having a free HNO₃ content of 3% by weight and a Bi content of11.1% by weight (=0.32 mol of Bi) were added at room temperature to theresulting solution. 106.23 g of solid Cr(NO₃)₃.9H₂O (=0.27 mol of Cr)were subsequently added. After heating to 60° C. and further stirring, ared solution (solution A) was obtained.

A further heatable 3 l stirred vessel made of glass was charged with2000 ml of water. 2.38 g of KOH (=0.042 mol of K) and 1124.86 g of(NH₄)₆Mo₇O₂₄.4H₂O (=6.37 mol of Mo) were subsequently added anddissolved at 60° C. The solution obtained was slightly turbid (solutionB).

Solution B was subsequently pumped into solution A, with the latterbeing stirred. 102.05 g of SiO₂ sol having an SiO₂ content of 50% byweight (“Ludox TM”, from DuPont) (=0.85 mol of Si) were added to theresulting dark yellow suspension at 60° C.

The suspension obtained was stirred at 60° C. for another 30 minutes andsubsequently spray dried (inlet temperature: 370° C., outlettemperature: 110–112° C.). The spray-dried powder obtained was admixedwith 4% by weight of graphite and subsequently tabletted to form solidpellets having a diameter of 5 mm and a height of 3 mm. The solidpellets were heated at 480° C. for 6 hours on a wire screen (meshopening: 3.5 mm) through which 100 l/h of air was passed in a mufflefurnace. The calcined pellets were broken up on a wire screen to givecatalyst granules K1 having an average granule diameter of 2–3 mm.

The oxydehydrogenation catalyst K1 had the nominal compositionMo₁₂Bi_(0.6)Fe₃Co₇Cr_(0.5)Si_(1.6)K_(0.08)O_(x).

Production of the Oxidation Catalyst K2

2500 ml of water were placed in a heatable 10 l stirred vessel made ofglass and 226.72 g of MoO₃ (from Fluka) (=1.575 mol of Mo), 25.17 g ofTiO₂ (“Finn Ti S 150”, from Kemira) (=0.315 mol of Ti), 6.99 g of Nb₂O₅(from H.C. Starck) (=0.052 mol of Nb) and 21.7 g of SnC₂O₄ (from Merck)(=0.105 mol of Sn) were added. The suspension obtained was refluxed for2 hours. After addition of 76.52 g of Sb₂O₃ (from Merck) (=0.52 mol ofSb), the suspension was refluxed for another 16 hours and subsequentlycooled to 50° C. 35 g of an aqueous Acronal solution having an Acronalcontent of 50% by weight (water-soluble polymer based on acrylic acid,from BASF) were then added, followed by addition of 140 g of formamide(from BASF).

To produce a coated catalyst, the resulting suspension was sprayed bymeans of a two-fluid nozzle onto 750 g of steatite spheres having a meandiameter of 2.5–3.2 mm (from Ceramtec). During the coating procedure,the steatite spheres were kept in motion in a coating drum which had aninternal diameter of 300 mm and was maintained at 150° C. and rotated at60 revolutions per minute. This resulted in a coated catalyst K2 havingan active composition content of 31% by weight.

The oxidation catalyst had the nominal compositionSbMo_(3.06)Ti_(0.6)Nb_(0.1)Sn_(0.2)O_(x).

Experimental Plant E1

The experimental plant was equipped with a feed unit and a reactor tubehaving two electrically operated heating zones which could be regulatedseparately. A tube reactor having electrically operated heating zonescan readily be used on a laboratory or pilot plant scale to replace ashell-and-tube reactor whose temperature is regulated by means of heattransfer medium circuits.

The reactor tube used had a length of 2 m and an internal diameter of 12mm. The feed was passed through the upright reactor tube from the bottomupward. The lower end of the reactor tube was provided with a mesh tosupport the beds. The reactor tube contained the following five bedzones, from the bottom upward:

Bed zone A (bottom): 15 cm of steatite spheres having a mean diameter of2–3 mm. Bed zone B: 80 cm of oxydehydrogenation catalyst K1. Bed zone C:30 cm od steatite spheres having a mean diameter of 2–3 mm. Bed zone D:65 cm of oxidation catalyst K2. Bed zone E (top): 10 cm steatite sphereshaving a mean diameter of 2–3 mm.

The bed zones A and B were located in the first heating zone and weremaintained at a temperature T₁, with the bed zone A functioning aspreheating zone for the oxydehydrogenation zone. The bed zones C, D andE were located essentially in the second heating zone and weremaintained at the temperature T₂, with the bed zone C (“inertintermediate bed”) functioning as preheating zone for the oxidation zoneand making it possible to precipitate relatively high molecular weightby-products from the first reaction zone.

The experimental plant was operated in a single pass. The reaction gasleaving the upper end of the reactor was analyzed by gas chromatography.

As feed gas stream, a mixture of 2% by volume of n-butenes (mixture of60% of 1-butene and 40% of 2-butenes) in air was fed in. The feed gaswas introduced at a rate of 140 standard l/h, which corresponds to aGHSV of 1600 standard l/[(l of catalyst K1).h]. The pressure at theupper reactor outlet was 0.01 MPa gauge pressure.

Example 1 (According to the Present Invention)

In example 1, the optimum reaction temperatures were set for bothreactions. The oxydehydrogenation over K1 was carried out at T₁=330° C.and the oxidation over K2 was carried out at T₂=400° C.

The values shown in table 2 were determined after a running-in time of 3days.

Example 2 (Comparative Example)

In example 2, the optimum temperature for the oxydehydrogenation over K1was set in both the reaction zones. T₁ and T₂ were thus 330° C. Thevalues shown in table 2 were determined after a running-in time of 3days.

Example 3 (Comparative Example)

In example 3, the optimum temperature for the oxidation over K2 was setin both the reaction zones. T₁ and T₂ were thus 400° C. The values shownin table 2 were determined after a running-in time of 3 days.

TABLE 2 Results from examples 1 to 3 Example T₁ [° C.] T₂ [° C.] C [%]Y_(MA) [%] 1  330 400 99 62 2* 330 330 99 32 3* 400 400 99 54*comparative example

Example 1 according to the present invention shows that optimumtemperatures in the two successive reaction zones, as can be realized ina shell-and-tube reactor having one heat transfer medium circuit in theregion of the first, feed-side reaction zone and a further heat transfermedium circuit in the region of the second, product-side reaction zone,results in a significantly higher yield of maleic anhydride than in thetwo comparative examples in which the temperature is the same in boththe reaction zones.

1. A process for preparing maleic anhydride by heterogeneously catalyzedgas-phase oxidation of an n-butene-containing hydrocarbon stream bymeans of oxygen-containing gases in a shell-and-tube reactor, whichcomprises conducting the gas-phase oxidation of the n-butene-containinghydrocarbon stream in a shell-and-tube reactor which has two successivereaction zones (a) and (b), (a) being the first, feed-side reaction zoneand containing at least one catalyst which is suitable foroxydehydrogenating n-butenes to obtain 1,3-butadiene, and (b) being thesecond, product-side reaction zone and containing at least one catalystwhich is suitable for oxidizing 1,3-butadiene to obtain maleicanhydride, and wherein the shell-and-tube reactor has, in the region ofeach of the reactions zones (a) and (b), at least one heat transfermedium circuit which is separate and independent from the heat transfermedium circuit of the other reaction zone.
 2. A process as claimed inclaim 1, wherein the reaction gas from the first, feedside reaction zoneis passed through an inert intermediate bed before it enters the second,product-side reaction zone.
 3. A process as claimed in claim 2, whereinthe inert intermediate bed used is a bed of inert material having anempty volume fraction of from 40 to 99.5%.
 4. A process as claimed inclaim 1, wherein the active composition of the catalyst suitable for theoxydehydrogenation of n-butenes to 1,3-butadiene is a multimetal oxideof the formula (I)Mo₁₂Bi_(a)Fe_(b)Co_(c)Ni_(d)Cr_(e)X¹ _(f)K_(g)O_(x)  (I), where thevariables have the following meanings: X¹=W, Sn, Mn, La, Ce, Ge, Ti, Zr,Hf, Nb, P, Si, Sb, Al, Cd and/or Mg; a=0.5 to 5; b=0 to 5; c=0 to 10;d=0 to 10; e=0 to 10; f=0 to 5; g=0 to 2; and x=a number determined bythe valence and abundance of the elements other than oxygen in (I).
 5. Aprocess as claimed in claim 1, wherein the active composition of thecatalyst suitable for the oxidation of 1,3-butadiene to maleic anhydrideis a multimetal oxide of the formula (II)SbMo_(a)Sn_(b)X² _(c)O_(x)  (II), where the variables have the followingmeanings: X²=V, Fe, Ni, Li, Ce, Nb, Ta, W, Ti, Zr, B, P, Al and/or Mg;a=1 to 10; b=0 to 5; c=0 to 3; and x=a number determined by the valenceand abundance of the elements other than oxygen in (II).
 6. A process asclaimed in claim 1, wherein the active composition of the catalystsuitable for the oxidation of 1,3-butadiene to maleic anhydride is amultimetal oxide of the formula (III)Mo₁₂V_(a)W_(b)X³ _(c)O_(x)  (III), where the variables have thefollowing meanings: X³=La, Mn, Fe, Cu, Al, Co, Ni, Bi, Ag, P, Zn, Cd,As, Cr, Sn, U, Ti, Nb, Ge, an alkali metal and/or an alkaline earthmetal; a=0.1 to 12; b=0 to 5; c=0 to 12; and x=a number determined bythe valence and abundance of the elements other than oxygen in (III). 7.A process as claimed in claim 1, wherein the oxydehydrogenation ofn-butenes to 1,3-butadiene in the first, feed-side reaction zone iscarried out at from 220 to 490°C. and the oxidation of 1,3-butadiene tomaleic anhydride in the second, product-side reaction zone is carriedout at from 190 to 500°C.
 8. A process as claimed in claim 1, whereinthe pressure at the reactor inlet is from 0.005 to 1 MPa gauge pressure.9. A process as claimed in claim 1, wherein the n-butene-containinghydrocarbon stream used is a raffinate II stream.